High velocity process for synthesis of organic compounds



June 2, 1953 H. e. M GRATH EIAL, 2,540,844

HIGH VELOCITY PROCESS FOR SYNTHESIS OF ORGANIC COMPOUNDS Filed Feb. 5. 1947 v 2 Sheets-Sheet 1 I 2.6 j 33 2+ 1 [as 24;

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HENRY 6 M fia/qr BY FTJM I 645 ATTORNEYS Patented June 2, 1953 UNITED STATES PATENT OFF ICE.

HIGH VELOCITYPROCESS FOR SYNTHESIS OF ORGANIC COMPOUNDS of Deiaware Application February 5, 1947, Serial No. 726,620

9 Claims.

This. invention relates to. thesynthesis of organic compounds; one aspect this invention relates to. thehydrogenation' of an oxide of carbon under conditions such that organic compounds. having more than one carbon atom per molecule are produced; In. another: aspect this invention relates to a method and apparatus for the synthesis of organic compounds. by the. hydrogenation; of carbonmonoxide with a relatively low ratio of hydrogen to carbon monoxide in the presence of a finely-divided powdered catalyst.

It l'l'asbeen known for some-time that hydrogen and. carbon monoxide may be made to react exothermicallyin the presence of certain catalysts and under specific: reaction conditions to form hydrocarbons and oxygenated organic compounds having. more than one; carboniatom' per molecule. In general, the synthesis of hydrocarbons hydrogenation of carbon menoxide is accomplished inthepresence of a metal or an oxide of a metalv chosen from group VIII of the periodic table as a catalyst at'pressures below' about 500' pounds per squarev inch gage and at temperatures below about 750 F.

Various-methods have been practiced to effect the reaction of hydrogen and carbon monoxide to produce organic compounds. Among these methods are those known as fixed-bed catalyst operations and fluid-bed. catalyst operation. The U fixed-bed. operation. comprises passing a, reaction mixture of hydrogen and carbon monoxide through a stationaiy bed of catalyst in. a reaction zonaand the fluid-bed operation comprises passing a reaction mixture through. a finelydivided' catalyst mass suspended in the reaction mixture in'the. reaction zone. characteristically, certain reactioncon'ditions. are. necessary for each 0t these processes. and for. the particular. catalyst used.

The synthesis feed gas or reaction mixture oomprises a mixture of about 1 to 2 molsof hydrogen per mol of carbon monoxide. and: may be prepared by the catalytic conversion of natural gas; steam, and carbon dioxide.

Themost recent development in the synthesis of organic compounds from hydrogen and carbon monoxide has been in the fluid-bed type operation; This type of operation" has had several apparent advantages over the fixed -be'd oper' ation. a-nd has yieldedorganic compounds of. high quality and in large. quantity- In such afiuidbed operation at a temperature of about 600 F. and at supenatmospheric pressures. using; a fluidized iron catalyst, a contraction of about 41 per cent to; about '20 per cent and. a carbon monoxide disappearance of about per cent toabout per cent have been observed. The selectivity of the reaction indicates that about 25 per cent to about 40. per cent of the CO is converted to C02. and oil-and water yields of about 100 to cc. per-cubic meter and about 60to. 12000. per cubic meter, respectively, are obtainable.

Even in view of the relativelygoodmesults obtained by the fluid-bed type operation certain inherent disadvantages have been found. In such i'luidebed operations inwhich the catalyst is, suspended in the reaction gas; classification of the. catalystv often occurs. There 15*,3150: a tendency for the fluid bed to settle after extended use oi the catalyst due to Wax and carbonaccumulation on the catalyst. These: tendencies have. required certain. design considerations in maintaining certain distances between the heat transfer surfaces. Furthermore, special consideration-must be made for the removal of heat liberated in the reaction, and with fluid-bed operation specially constructed apparatus is necessary for the removal of such heat with attention directed to the fluidized catalyst itself. It is much to be desired; therefore, to provide a process and apparatus which overcome these difficulties, at least partially.

It is anobject of this invention to provide a process for thesynthesis of. organic compounds having, more than. one carbon atom permole.- cule.

It is anotherv object of this invention to produce hydrocarbons and oxygenated compounds by the reactionv of. carbon monoxideand'hydrogen in'the presenceof a hydrogenation catalyst;

Another: object. of. this invention is: to provide an improvement in the synthesis of. hydrocarbons in the presence of a finely-divided: fluidized catalyst. in the reaction mixture.

Still a further object is. to. provide a method for the synthesis of hydrocarbons using a rela.- tivelylow feed ratio of' hydrogento carbon men-.- oxide.

Still another object is to provide a fluidized process for the hydrogenation of carbon monoxide in which the catalyst life is extended and prolonged.

Various other objects and advantages will become apparent to those skilled in the art from the accompanying description and disclosure.

Much to our surprise, we have found that an oxide of carbon may be hydrogenated to yield organic compounds having more than one carbon atom per molecule in the presence of a finely-divided hydrogenation catalyst which is present in the gaseous reaction mixture in a much smaller amount than heretofore thought possible. We have found that a catalyst concentration in the gaseous reaction mixture of less than about 18 pounds per cubic foot of gas is adequate to carry out the reaction between hydrogen and carbon monoxide with a comparable yield of products to other synthesis processes using a much larger concentration. Furthermore, the contact time between catalyst and reactants in the reaction zone is less than seconds per pass, usually between about 0.5 and about 3 seconds. Contact time is defined as the time of contact of the synthesis gas with catalyst per pass. To effect the reaction in the presence of a finely-divided iron catalyst present in the reaction mixture in the amount heretofore described, according to an embodiment of this invention a reaction mixture of hydrogen and carbon monoxide is passed upwardly through an elongated, substantially vertical, reaction zone at a velocity greater than 8 feet per second and as high as 40 feet per second, although velocities as low as 5 or 6 feet per second may be used under certain conditions without departing from the scope of this invention.

At such velocities, the finely-divided hydrogenating catalyst, such as iron, is entrained or suspended in the reaction mixture in an amount between about 1 and about 18 pounds per cubic foot of gas and forms a continuous catalyst phase in the reaction zone. In some cases a concentration as high as 25 pounds per cubic foot is desirable but preferably lower concentrations are used. The reaction mixture and the catalyst which is entrained in the flowing gases are passed through the elongated reaction zone and are removed from the upper portion thereof together after less than about 5 seconds of residence therein. According to this invention, the catalyst is entrained in the reaction mixture in the reaction zone and flows therethrough in the direction of flow of the gases under conditions such that the conventional dense, pseudo-liquid catalyst phase is not formed.

In conventional fluid-type operations the catalyst forms a so-called dense phase catalyst bed in the reaction zone and consequently remains largely in the reaction zone itself until removed. Actually, two phases are formed in the reaction zone, a dense catalyst phase in the lower portion and a dilute phase having only a small amount of catalyst in the upper portion. The concentration of catalyst in such a dense phase is at least 25 pounds or 35 pounds per cubic foot of gas and usually between 50 pounds and 120 pounds per cubic foot.

It has been found in operating with a continuous catalyst phase according to this invention that a hydrogen to carbon monoxide feed ratio as low as 0.7 to 1 can be used with success. This ratio will vary from about 0.7 to 1 to about 1.5 to 1 and higher ratios may be used, if desired,-

4 but in any event the ratio is less than that generally required in conventional fluid-bed operations.

The catalyst employed in the present invention is a finely-divided powdered catalyst of a metal or metal oxide which is or becomes in the reaction zone a catalyst for the hydrogenating reaction. Finely-divided metallic iron or iron oxide or a mixture of metallic iron and iron oxide are an example of the catalyst employed in this invention. Preferably, a metallic iron catalyst is used in the finely-divided form. Other metals and metallic oxides may be employed which are effective in catalyzing the hydrogenation of carbon monoxide, such as cobalt, nickel, and other metals of group VIII of the periodic table. While the catalyst powder usually consists of such catalytic metals or their oxides, it may also include a minor amount of promoting ingredients, such as alkalies, alumina, silica, titania, thoria, manganese oxide, and magnesia. Also, the catalyst may be supported on a suitable support, such as a bentonite type clay, silica gel, Super Filtrol and mixtures of these supports. In the following description, catalyst powders consisting of a metal and/or a metal oxide and containing at most a minor proportion of promoters are referred to as finely-divided metal catalyst.

The exact chemical condition of the catalyst in its most active form is not certain. It may be that the active form is present when the metal is at an optimum degree of oxidation and/ or carburization; consequently, a metallic iron catalyst which is in a reduced condition when first contacted with the reactants may reach its state of highest activity through being oxidized and/or carburized in the reaction zone. Therefore, in this specification and claims, the catalyst employed is described by reference to its chemical composition when first contacted with the reactants.

The catalyst is employed in a fine state of subdivision. Preferably, the powdered catalyst initially contains no more than a minor proportion by weight of material whose average particle diameter is greater than 250 microns. The greater proportion of the catalyst mass, preferably, comprises a material whose average particle diameter is smaller than microns including at least 25 weight per cent of the material in a particle size smaller than 40 microns. An example of a desirable powdered catalyst is one which comprises at least '75 per cent by weight of material smaller than microns and at least 25 per cent by weight of materials smaller than 40 microns.

The temperature of reaction for the hydrogenation of carbon monoxide is generally between about 300" F. and about 750 F. With a metallic iron catalyst, temperatures between 450 F. and 750 F. are usually employed. With a cobalt catalyst usually a temperature below 540 F. is sufficient for the hydrogenating reaction. Generally, the pressures employed are somewhat above atmospheric and range from about 10 pounds to as much as 500 pounds per square inch gage, preferably between about 80 pounds and about 300 pounds per square inch gage.

In effecting the reaction it may often become necessary to cool the reaction zone to maintain a relatively constant temperature. Various methods of cooling the reaction zone itself, such as by external cooling means or by injection of a cooling medium directly into the reaction mix- 5 ture, may bepracticedwithout? departing from: the scope-of this-invention. Furthermore,- itmay.= often become necessary to preheatthereactiom mixture priorto entry into. the reaction zone, andalso the-catalyst may :be preheated' before introduction into the reaction mixture. However; thecooling and" preheating are factorswhich. Will be: characteristic of the particular apparaetusibeing used" andthe particular;condition una der which the reactionis effected.

Generally the. reaction. zone. itself will. comaprise. a single or." a: multiple number; of: conduitsc or. tubes. of an. inside diameter between. about; 1 inch:.-.and:. about 6 .-inches.z. Preferably; the -diame eter of? the reactiontubes is=between about 1'. inch andlabout,;2;5 inches... Itzis knowrrthat the diameter of the tube isnof: considerable:importance: when entraining a catalysttinthe-reaction: mixe tureI since. the. Wall efiectl. of i the: tube; itself; has: alconsiderabl'e; efiect on; the; disposition; of the catalyst in the I'GEClJi'OIICxStI'ElIILp Consequently,-, it has? been found? that; generally? the; reaction tube :size-shouid'ibe.=.belowt-abouts.2a5: inches in adj-r ameter. for optimum ;operations.- However, with certain; catalysts-v and. with. certain conditions; of.;op eration;c tubes,.:r:reactionzones: much ;-la-rger;- in;.diameter may, beused; From theastandpoint of; cooling; the; reaction zone,. smaller: tubes: are.

also: desirable": since: they present larger heat transfer surfaces;-

According to; a preferred; embodiment. of i this. invention, a: fresh feedgashaving alhydrogen to carbonzmonoxidaratio: hi her: than the ratio inwhich these compounds; are; converted; tov other;

compounds islemployedz'and'ztherratio. of hydrogen in ,ap-volumetric ratioa of recycle to. fresh feedgasz: of about 0.55.21 to about:10:1,.generallyyabout1:1L The ration-S. hydrogen. to: carbon-.zm-oncxide in'lthefreaction zone itself: iss; usually abouti 1 to about -3: 1 and z according to:

to about: 5:1. or.- 61:1:

this process maybemaintainecl at: about '1 1 with out detrimental effect on the synthesis reaction; Theiratioof hydrogen to carbon. monoxidein: the fresh" feed itself I may be. considerably lower than in the reaction zone and usually: ranges from about-Dilute? about 115:1. For very'lowfeed. ratios of. less than 0.5. :1: the -ratio in the reaction; zone Titself-maybe -evenless than the feed ratio as. av result of the consumption ofhydrogen rel ative-tocarbon monoxide being greater than the feed ratio. Once'through operations-Without-recycle, although not generally useclgare within the scopetof. this invention.

The linear velo city'of the gaseousv reaction mixture passing upwardly through-thereaction: zone is conveniently expressed in' terms of superficial velocity, which i is: the linear velocity the feed stream would assume if "passed throughthe reactor in the absence.of catalyst, an'd takes into account: the shrinkage-in volume: caused by the hydrogenation reaction: As previously stated;

these superficial velocities are aboveabout 8 feet per second,- preferably above about-12. feet-pe1" second, and-quay be as high as"40"feet per second or---higher without departing from thescope of this invention. Thesuperficiallinear velocity is calculated'from. thearithmetic average of 'thegas: rate at the bottom and top of the-reaction zone; The latter is arrived at by correcting the outlet. gas volumefor water and: hydrocarbons con.- densed in the receivers, with correctionsf or pressure and; average: catalyst temperature. Con-- tact times are calculated by dividing the length: of the reaction zone by the superficial velocity.

The concentration-bf the catalyst in: the gaseous reaction mixture in the' reaction zone is: usualli less than about: 18 pounds per'cubic foot. of gas; and? preferably between about 3" pounds and about 12* pounds-per cubic foot; The actual; concentration required in the-aboVe-range' will depend to a cert'aimextent-upon the amount of" inert g-as in' thev reaction zone and also upon the?- accumulationi of 5 carbon and wax on v the catalyst particles as the operation proceeds-z The accu mulation oi wax and carbon" on the catalyst-decreasesthe weight of catalyst" per" cubic f'oot of gas thus the above values represent'theextreme limits and may vary in accordance with this dismission:-

Although the-invention has been described with reference 'to an upwardly-flowing gaseous stream: of reactant's and catalyst, it shouldbe unclerstoc d thatthe catalyst and reactants may fi'ow together downwardly, horizontally; or evenangularly, through 1 a reaction zone without departing fromthe scope ofthis invention. Iinportantrequire' ments of this invention are that the weight of the catalystper cubic foot'of g-a's andthe'contact timebewithin the limits previously'discussed It has been found that by upward fibwingof gas-- through a substantially vertical reaction zone the Weight of "catalystpei" cubic foot of F gas and the reaction-time can be controllediconveniently and accurately and ion onereason is the preferable method of operation.

In operating a synthesis= process according tothis' invention with an 'iron'catalystand'at a: temperature b'etween about'550 F; and about 650 h. at relatively low super-atmospheric pressures -a= contraction ofa-bout 5 to about i per cent has beenobserved." The carbon'monoxidedisappear anceis ab'out "10 percentto about 88per cent andtlie selectivity of the "reaction illustrated' by the conversion ofcarbon monoxide to carbon dies-'- ideis about-w per cent to'a'bout 30 1661" cent;

Condensed oil and-water yields of' 'about- 30-to about I and ab'outfie to about cc. per cub'ic' meter of fresh feed gas, respectively, are=ob-- tained by operating accordingto the presentprocess and may contain I appreciable quantities of organic chemicals;

Certainparticular advantages of the present. process have=been' observed; One ofthese ad vantages whichis ofgreat importance is the-fact that a relatively low-hydrogen to carbon monoxido -feed ratio can be utilized I andasa result a proportionate decrease in the cost of the production of hydrogen as efiected. At theextremely high velocity; capable of being usedby the pres ent-processiand apparatus much less catalyst isused andmuch-greater efiiciency of heattransfer'is obtained. Furthermore, temperature conditions can be easily and readily controlled at such high velocities and at suchsh'ort contact times. By theme of a-short contact time higher 1 selectivitycan be obtained as the resultof decreased sid'e'reacti'ons; suchas over-polymerization; Less'th'an 3f per-cent pjer week'wax accumulation on the "catalysthas "been" observed with the present method which compares with as much as 23 per cent per week wax accumulation with a low velocity fluid-bed type operation. Carbon deposits upon the catalyst have also been observed to be much less than those observed in low velocity operations. A carbon deposit of less than about 2 per cent per week for the present process as compared to above about per cent per week for the low velocity process has also been observed.

As a result of high velocity and increased heat transfer, it is possible to use much greater temperature gradients between the catalyst and a cooling medium, if one is used, with the present process. In the low velocity fluid-bed operation a temperature difference between coolant and reaction mixture between F. and 100 F. is conventional. However, with the present process a gradient considerably in excess of 100 F. is possible. Also, relatively cold gas feed may be employed to the reaction zone. Mixing is so rapid under such conditions that no deleterious effect is observed.

The invention Will be described further by ref erence to the accompanying drawings which are views in elevation, partly in cross-section, of suitable apparatus for carrying out the process of the present invention. Fig. 1 of the drawings is an elevational view diagrammatically illustrating a reaction zone and suit-able auxiliary equipment for carrying out one embodiment of the present invention. Figures 2 and 3 of the drawings are other reaction chambers embodying the essential features of the present invention and may be substituted for the reaction chamber shown in Fig. 1 of the drawings.

In Fig. 1 of the drawings a synthesis gas comprising hydrogen and carbon monoxide present in a ratio between about 0.711 and about 1.421 is obtained from any suitable source. For example, a suitable source of hydrogen and carbon monoxide is the conversion of steam, carbon dioxide, and methane in the presence of a suitable catalyst, such as nickel. The resulting mixture of such a conversion usually contains sulphur and sulphur compounds, and the gas is preferably purified to remove such compounds therefrom. If a sulphur resistant catalyst is used the purification step is unnecessary. After purification in conventional manner known to those skilled in the art, the mixture of hydrogen and carbon monoxide is introduced into the lower end of a 26 foot length 1 inch diameter conduit or tubing 8 of Fig. 1. Conduit 8 is a curved conduit which has a major portion thereof positioned substantially vertically. Conduit 8 is also lagged. The gaseous reaction mixture is passed upwardly through conduit 8 and catalyst from a standpipe I0 is introduced into the flowing gaseous stream in the lower portion of conduit 8, as shown. The velocity of the gas in conduit 8 is maintained above about 8 feet per second in the vertical section in order to prevent the formation of a pseudo-liquid dense phase of catalyst in the vertical section of the conduit, but instead to form a continuous catalyst phase of relatively dilute concentration. Usually the velocity of the gaseous stream in conduit 8 is between about 12 and about 40 feet per second and at such velocities the catalyst is entrained in the gaseous stream and passes overhead into the upper portion of standpipe I0 with the gaseous stream itself. Since a continuous catalyst phase is present in conduit 8, the amount of catalyst taken overhead into the standpipe is approximately equivalent to the amount of catalyst introduced into the lower portion of conduit 8 from standpipe l0.

standpipe I0 comprises two substantially vertical, concentric, extra heavy steel pipes; an outside pipe II and an inside tubing I0. The outside pipe II is welded at its ends to the inside tubing I0 to form an enclosing jacket which may be filled with a liquid as a cooling or heating medium. The cooling or heating medium i introduced into the annular space formed by the two concentric pipes through line I2 and may be withdrawn therefrom through line I5. In some instances the cooling or heating medium may be introduced through line I5 and removed through line I2, if desired. In another embodiment where the liquid introduced into the annular space is evaporated therein and the latent heat of evaporation is used to cool the catalyst, liquid is introduced through line I2 and vapors are removed, also through line I2.

Catalyst passes from conduit 8 into a conical section I4 which has a larger diameter than conduit 8 and thereby the velocity of the gases are diminished and the catalyst separates from the gaseous stream and flows downwardly into conduit ID. A slide valve I3 in the lower portion of standpipe regulates the flow of catalyst from conduit I0 into conduit 8. The upper end of conduit I0 is connected by means of a conical section I4 to an enlarged conduit I6. Conduit I5 facilitates the disengagement of the catalyst from the gas stream after the passage of the latter into conical section I4. Conduit I6 is connected by means of manifold 11 with conduits I8 and I9. Conduits I8 and I9 contain ceramic filters 20 and 2I which are constructed of porous material, such as alundum, permeable to the gases and vapors emerged from conduit 8 but impermeable to catalyst fines entrained in the gaseous effluent. Filters 20 and 2| are cylindrical and are closed at the bottom ends. A substantial annular space is provided between the wall of the filters and the wall of the enclosing conduit for the passage of gases and vapors and entrained catalyst upwardly through the annular space between the filters and conduits I8 and I9. The upper ends of filters 20 and 2I are mounted inside conduits I8 and I8 by means of enclosure means 22 and 23. The gases and vapors must pass through either or both filters 20 and filter 2| to reach outlet conduits 24 and 26.

Union of the various conduits in the upper portion of the standpipe I0 is made by welding.

When using a metallic iron catalyst in a finelydivided state, usually between about 40 and about 150 microns, the temperature of reaction in conduit 8 is between 550 F. and about 650 F. Apressure of about 80 pounds per square inch gage has been found to be quite satisfactory. However, various pressures above and below this may be used without departing from the scope of this invention. With a velocity greater than 10 feet per second in conduit 8 the reaction time is less than 3 in the vertical section of conduit 8. Similarly, if

the loading is cut to about 4 pounds per cubic foot of gas the concentration of catalyst in the vertical section of conduit 8 is about 8 pounds per cubic foot. In all cases the velocity of the gas passing.

:9 through conduit 8 is maintainedtabove about 5 or 6 feet per second in order to prevent a formation of a dense pseudo-liquid phase of catalyst and to assure a continuous phase of catalyst in conduit 8. In operating at such high velocities the catalyst is entrained in the gaseous mixture and flows from the lower portion of conduit 8 to the upper portion thereof and settles or separates from the gaseous mixture in conical section It and enlarged conduit [6. The gaseous portion of the mixture from conduit 8 flows upwardly through the filters 20 andZl and into the respective outlet conduits 24 and 26. Separated catalyst flows downwardly through conduit Ill and by the regulation of valve I3 is introduced into conduit 8 at the desired rate. The density of the catalyst in conduit I0 is usually about 50 pounds to .110 pounds per cubic foot of volume.

In some instances, the heat of reaction may be removed by cooling the catalyst and using the sensible heat of the catalyst as a means for cooling the reaction mixture in the reaction zone. To accomplish this end, the catalyst in this particular apparatus may be cooled in conduit Ill by introducing a liquid, such as water orDowtherm, through conduit l2 into the annular space between concentric conduits l I] and H. The evaporation of the water or Dowtherm in the annular space removes a large portion of the heat in the catalyst. The cooled catalyst is then introduced into conduit 8 forrecycling through the reaction zone or conduit8. Conduit 8 may be cooled directly itself by indirect heat exchange (not shown) without departing from the scope of this invention. Various other methods known to those skilled in the art may be used to cool either the reaction mixture in conduit 8 or the catalyst in conduit l0 without departing from the scope of this invention.

Since the pressure differential between just below slide valve 13 and the upper portion of reaction zone 8 may vary to a considerable extent, it is necessary to control slide valve 13 to compensate for the pressure differential in order to obtain a constant flow of catalyst from standpipe l ll into conduit 8. This control of slide valve '13 is obtained by connecting adifferential pressure recorder 28 by means'of conduits .27 and 29 to the upper and lower portions of standpipe 10, as shown, and transmitting changes in pressure differential to slide valve l 3 by means-of element 3! .so that when the pressure differential is creased valve is is closed slightly and when the pressure difierential is decreased valve 13 is opened slightly. The concentration of the catalyst per cubic foot of gas in the vertical section of conduit 8 may be determined by connecting a pressure differential recorder (not shown) on the vertical section of conduit 8 and calibrating the recorder readings in terms of concentration of catalyst. The gaseous effluent from either conduit it or la is passed through'filters "20 and-2i into outlet conduit 24 and '26, respectively. Usually only one outlet conduit is used at atime. Thus, for example, the gaseous effluent passes through outlet conduit 2a to conduit through condenser 34 where the effluent is'cooled to about 40 F. at operating pressure and then passed to accumulator 3.5. In accumulator :36, gaseous components are separated from liquid components of the cooled eliiuent. Uncondensed cor-npenents of the eliluent, such as hydrogen, carbon monoxide, methane, propylene, 'butylene, light naphtha, and organic oxygenatedcompoun'ds, are recycled by means of conduits 3'8 and 44 and a and carbon monoxide.

compressor or blower 45 to the lower portion of conduit 8 in a ratio of about 1:1 to about 5 or 6:1 of volumes of recycle to volumes of fresh synthesis gas. The amount of unreacted hydrogen and carbon monoxide in the recycle gas determines how much the ratio of carbon monoxide and hydrogen in the reaction zone itself will deviate from the ratio in the original feed. As shown, the recycle gas is introduced into the fresh feed before the catalyst is introduced into the gaseous mixture, however, the recycle may be introduced into the gaseous mixture after the catalyst is introduced into the feed stream or the feed gas may be introduced into the gaseous mixture after the introduction of the catalyst without departing from the scope of this invention.

After passage of the gaseous effluent through filter 29 for a time, the filter becomes coated and clogged with catalyst fines which have not settled out from the gaseous efiiuent. In order to remove these fines from the catalyst filters so as to ensure continuous passage of the gaseous effluent through the filters and so as to recover the cata- .lyst, the course of gaseous eiiiuent is changed to flow through filter 2| and conduit 26 and a portion of the uncondensed efliuent is passed from accumulator 36 by means of conduit 138, com pressor 4i and conduit 24 to filter so. The pressure of the gase blows the fines from the filter into conduit l8. The fines then settle in conduit l8 tostan-dpipe I fl. Other gases than the uncondensed efiiuent may be used to remove the fines from the filter and may be introduced through line 42, if desired.

Liquid condensate in accumulator 36 is passed through conduit 39 to separating means 18, which may represent various separating units, such as distillation columns, absorption units, extraction units, and the like. In separating means 48, water is separated from organic compounds and removed through line 49; oxygenated organic compounds are separated therein and removed through line 5!; and hydrocarbons are separated and removed through line 52.

Uncondensed gas from accumulator 36 not used for recycle, etc, is removed from the system through conduit til and passed to oil and chemical recovery equipment not shown.

It has been found that operating a synthesis process according to this invention, in which the synthesis gas is passed through a reaction zone at a, high velocity, good yields of products are realized. Ordinarily one would believe that insuflicient catalyst for accomplishing the desired reaction would be carried by the gas at such high velocities, but it has been found that within the range indicated, suflicient catalyst is carried by the to effect the reaction between hydrogen If desired, the synthesis entering conduit 8 may be preheated; but it has been found that preheating the gas is unnecessary in most instances and that the contact of thehot catalyst from standpipe It with fresh fee'dgases entering conduit 8 does not cause balling or agglomeration of the catalyst mass. It has. also been found, as previously mentioned, that the wax and carbon content of the catalyst with extended use is much less than that observed in theconventional fluid-bed operations, and it is believed the reason for this is that catalyst eddy-currents are minimized in high velocity operations which greatly shortens the contact time between catalyst and reactants per pass. Longer catalyst life may also result from shorter 11 contact time per se between reactants and catalysts. The above theory is offered merely as a possible explanation of the extended life of the catalyst realized in the present invention and is not considered to unnecessarily limit the invention.

Fig. 2 of the drawings is another arrangement of apparatus suitable for carrying out this invention. The apparatus shown in Fig. 2 may be substituted for conduit 8 and standpipe I of Fig. l. The filter sections of Figures 1 and 2 are the same. Accordingly, a synthesis gas comprising hydrogen and carbon monoxide is passed into a reaction zone I0 through a line 68. Reaction zone I0 comprises a bundle of substantially vertical tubes "II through which the gases pass upwardly into a cylindrical section I4 of a larger total cross-sectional area than the total crosssectional area of tubes II. The outer surface of tubes II are sealed off to form an annular space 12 between the inner surface of the outer shell of reaction zone It! and the outer surface of tubes II. A cooling (or heating) fluid may be passed through annular space I2 to cool (or heat) the reaction mixture as it passes upwardly through tubes 'II. Tubes II correspond to conduit 8 of Fig. 1. Upon reaching enlarged section 14 the velocity of the gaseous mixture is decreased to such an extent that the catalyst settles from the efiluent and passes down through a standpipe I3 into the lower portion of reaction zone where the catalyst falls or is drawn into a high velocity gaseous stream passing upwardly into tubes II. The lower portion of reaction zone I0 and tubes II are of such a cross-sectional area that the catalyst is entrained in the gaseous stream.

The cooling medium is introduced into the annular space I2 through line 16 and is withdrawn therefrom through line 11.

The reaction effluent from reactor 10 passes upwardly through conduits I'I, I8, and I9 of which the latter two contain filters as previously discussed with reference to Fig. 1. Conduits I'I, I8, and I9 are the same as conduits IT, IS, and I9 of Fig. 1. duits I8 and I9 through outlet conduits 24 and 26 to conduit 33. The efiluent in conduit 33 of Fig. 2 is condensed and separated according to the description of 1 and a portion of the uncondensed gases may be recycled (not shown) to conduit 68, if desired.

The cross-sectional area of conduits 'II with respect to the quantity of gases flowing therethrough is such that the velocity of the gases is greater than about 8 feet per second, while the cross-sectional area of enlarged section I4 is such that the velocity of the gases is below about 5 feet per second so that the catalyst may settle from the gaseous efiluent. In section I4, the catalyst may form a dense pseudo-liquid catalyst phase having a concentration of catalyst greater than about pounds or pounds per cubic foot of gas. In such a case, reactor 10 has two reaction sections, tubes II in which the gas flows at a relatively high velocity and with a relatively low concentration of catalyst and enlarged section T4 in which the gas flows at a relatively low velocity and with a relatively high concentration of catalyst. The same effect may be achieved in the apparatus of Fig. l of the drawings by extending the length of section I6 and adjusting the cross section thereof such that the velocity of the gases therein are appropriate for the formation of a dense pseudo-liquid catalyst phase. In

The gaseous cfiluent passes from conthis manner, a high velocity continuous catalyst phase will exist in conduit 8 and a dense pseudoliquid catalyst phase will exist in section I6.

A moveable valve 69 is provided at the lower end of standpipe I3 to control the flow and dispersion of the catalyst into the gaseous stream. Valve 69 may cause an aspiration efiect by the deflection of the gases passing by and as a result of which the catalyst is drawn into the gaseous stream.

Fig. 3 gives a diagrammatic illustration in elevation of another arrangement of apparatus for th synthesis of hydrocarbons according to the present invention. The apparatus in Fig. 3 is very similar in operation to the apparatus of Fig. 1 and Fig. 2 and thus only a brief discussion of its operation will be included. A synthesis gas passes through a conduit I04 to a reaction chamber I06. Reaction chamber I06 comprises a bundle of reaction tubes I01 surrounded by a shell I08 to form an annular space I09 between the outside diameter of tubes I01 and shell I08. Annular space I09 is for the circulation of cooling fluid around reaction tubes I 0'! in order to maintain the temperature of reaction substantially constant. The cooling fluid, such as Dowtherm, enters annular space I09 through conduit II3. The pressure in annular space I09 is such that the Dowtherm boils below the desired temperature of reaction in tubes I01. The vaporized Dowtherm passes from annular space I 09 through conduit II2 to an accumulator II4. From accumulator I I 4 vapors pass through line I I5 to a con denser I I 6 in which substantially all of the Dowtherm vapor is condensed. Condensate from condenser I I6 passes to accumulator I I4 through conduit I i8. From accumulator II4 liquid Dowtherm is recycled to annular space I09 through conduit II3. Any uncondensable vapors are removed from the system through conduit I I1.

A reaction eflluent comprising reaction products and finely-divided catalyst entrained in the reaction efiluent is passed to a settling and accumulation chamber I20 through conduit III. Chamber I20 comprises an upper settling chamber I2I and a lower accumulation chamber I28. The cross-sectional area of settling chamber I2I is such that substantially all of the catalyst separates from the gaseous effluent and flows downwardly through a funnel-shaped septum I 21 into the lower portion of accumulation chamber I28. Gases containing a small amount of fine catalyst pass upwardly in settling chamber I2I into a cyclone separator I 22. Gases from cyclone separator I22 are removed therefrom by conduit I23 and may be treated in the manner heretofore described. Separated fine catalyst is removed from cyclone separator I22 by means of conduit I 24 and passed to funnel-shaped septum I21. as shown. Finely-divided catalyst accumulates in accumulation chamber I 28 to a level above the end of the funnel I2'I in order to prevent the passage of the gaseous ellluent downwardly through funnel I21. An aeration gas, such as hydrogen or recycle gas, is introduced into the lower portion of accumulation chamber I28 through conduit I32 and is injected into the accumulated catalyst therein by means of dispersion means I33 which may comprise a perforated conduit or the like. Accumulator I28 may be maintained at a substantially higher pressure than settling chamber I2I by introducing a gas therein through conduit I3I, if desired. By maintaining the pressure higher in chamber I28 than in chamber I 2I passage of gaseous eflluent Microns:

'used, such as'the filter means of Figures 1 and Various other modificationsoi Fig. 3 may become obvious to those skilled in the art without departing from the scope of this invention Assuming a reactor inlet gas volume of 25,000 standard cubic feet per hour, a reactor consisting of four tubes having an inside diameter of 2 inches would operate at a maximum. linear velocity of reat per second. Using'a tube length of 36 feet, a velocity of 9 feet per second would correspond to about 4 seconds of contact time per pass, allowing for contraction. A reactorcapable of a 40 feet per second linear velocity and 5 seconds contact timeper pass would involvea flow path of about 200 feet in length. A single tube having a2 inch inside'diametermight be substituted .for the four tubes havin a 1-inch inside diameter with substantially thesame velocity and throughput.

Various .modifications and alterations of the apparatus. and iiow shown inliigures 1., 2, and.3 may be practicedby those skilled in the art without departing from the scopeof this invention.

Various coolers, condensers,.distillation units, and other separating means have not been shown-for a matter of convenience andsimplicity, but their presence will be obvious to those skilledin the art.

The following example isofieredas a means of better understanding the application. of the present invention to the hydrogenation of carbon monoxide and the specific ,recitationof certain limitations therein is not considered unnecessarily limiting to the present invention.

EX-ANIPLE In this example an, iron catalyst was employed, which iron catalyst was preparedby fusing-and subsequently reducing Alan Wood .ore. :The catalyst contains metallic iron, alumina titanium dioxide, and silica, and about 1.2 .to about 1.4: weight per cent potassiumcalculated as the oxide. The potassium. oxide is incorporated with the catalyst prior to fusion. The particle.size'ofthe catalyst employed .in the runsofthis example is shown in Table I below:

Table I rownssso IRON CATALYSTPARTICLE SIZE Roller An alysis Recovery, per Density (be-sis water) Having prepared ,a catalyst of the -;desired properties and-the (required size, the catalyst was "tion' zone. from the reaction zone aftera relatively short introduced into a conduit, similar to conduit 8 of Fig. 1 of the drawings, in which a mixture of hydrogen and carbon monoxide was flowing upwardly at a relatively high velocity. The conditions of reaction and the analysis of the product is shown inTable II below. .In substantiallypallrespects the apparatus used for obtaining the data shown in Table lI-was the same as the apparatus of Fig. 1 of the drawings. In obtaining 1 data various gas velocities as well as concentrations of catalyst in the reaction zone (conduit 8) were usedto determine their effect on the reactionand product. Also it will be noted that the pressures used varied from -,to pounds per squareinch gage. .In some 1runs, especially the 1ater runs, :the-temperature of the inlet feed gas was decreased to a relatively low value to determine the effect of contact between relatively hot catalyst-and relativelycold gas and whether tempera- ;ture conditions could be maintained. inathe .reac- The reaction efiluent was withdrawn contact time with the catalyst. The catalyst iseparated byggravity from the efliuent in an enlarged section of the apparatus and passed by means of astandpipe, such as .conduit ldof Fig. 1, to the point of introduction of .thecatalystinto the'synthesis'gas stream. The efiluent passed through filters to separate fine catalyst therefrom. The filters were cleaned intermittently .or continuously by flowing recycle gas back through themas previously explained with referenceto Fig. 1 of the drawing. The eiiiuentthenpassed through a conidenser at 40 F. andat operating pressure and. uncondensedgases were recycledcto a point just :before the first contact-oi -:catalyst and synthesis gas.

Catalyst in. the standpipewas maintained at about-600 F. bymeans of electricalheatingi-elemerits wound aroundthe l inch jacket surrounding the 2 inch standpipe. .Variouspredetermined loadings of thecatalyst into the synthesis gas stream wereused and the itendencyfor variation .in loadingsior any-particular run caused by vari- .ation in differential pressure inthe reaction zone was minimized by controlling aslide .valveon the bottom of the standpipe by a conventional'diiier- .entiaLpressure recorder responsive to-the differential pressure between the top and .bottomof the standpipe. The. differential pressure recorder was -set at various readings'of inches of water depending upon the catalyst loading desired. The concentration of catalyst in the reaction'sone or elongated conduit was determined by thedifferential pressure 15 foot verticalsection of the conduit. The overall lengthof the reaction lzone was about Zdfeet and .3 /2inches, i. e., fron1 the point ,thecatalyst was first contacted with the synthesis gas and the point where the catalyst was separated from the reaction efiluent.

"The velocity ofthe gas stream so great that the catalyst was entrained in the synthesis gas stream 'throughoutithe reaction zone in.a continuouspha'se without the formation of the conventional dense pseudo-liquid phaseof catalyst.

'Gas'bleeds into the instrument lines Were used to preventcatalyst from clogging instrument lines andjinstruments. .In some cases recycle gaswas used as thebleed gas, in other instances hydrogen onfresh feed gas were used.

The catalyst in the standpipe was aerated with recycle :gas or combined inletgas in most instances-however othergascs, such as carbon-dioxide, hydrogen, nitrogen, and steam could have been. used .ifdesired.

Table II Run No 1 3 4 5 6 7 9 12 13 14 15 18 19 20 21 24 25 Hours on run 6 6 16 18 18 150 12 24 22 6 18 18 24 24 24 48 24 Operating conditions:

Pressure, 1). S. i 80 80 80 80 80 150 150 150 1.50 150 150 150 150 150 150 150 150 Ratio of recycle to fresh feed 6.0 7. 7 1. 7 1. 8 6.2 4. 92 4. 7 8. 5 6. 5 8.1 6. 6 6. 5 6. 4 5. 6 5. 4 6. 6 5. 8 Fresh feed, 0. FJH 143 150 155 191 251 224 106 133 105 127 135 128 145 148 126 171 Gas ratio, H2:C0

Fresh feed 3 3 3 3 3+ 3 2 1.4 1.4 1.4 1.4 1.4 1.4 1.4 1.4 1.4 1.4 Inlet to reaction zone 4.5 14 4.9 3.9 2.0 2.2 2.0 2.2 2.6 2.6 1.7 2.0 2.1 standpipe:

Temp, F 593 598 598 596 605 597 597 590 597 601 600 604 508 600 595 559 591 Density, P. C. F. 97 100 103 99 102 106 87 87 82 80 78 69 67 67 66 55 49 Lin. velocity, F. P. 0.4 0.4 0.10 0.17 0.18 0.14 0.13 0.11 0.11 0.08 0.10 0.10 0. 09 0.09 0.11 0.10 0.12 Reaction zone:

Temp. F., in zone 585 595 594 596 603 601 600 600 604 604 602 602 595 500 585 540 602 Gas inlet temp, F 572 616 604 603 605 576 603 600 600 600 600 550 500 450 400 150 615 Line velocity, F. P. S.

(superficial) 15 19 7 8 15 14.3 13. 4 10. 5 10.4 10.2 10.3 10.6 9. 8 8.9 9. 5 8. 5 10. 1 Catalyst loading (inches water 53 35 49 70 70 70 59 58 59 58 57 Concentration 01 catalyst, P. C. F a, 3.5 5.9 3.5 3.0 4.9 8.4 8.0 8.7 9.4 10.1 11.2 8.2 8.4 8. 7 8.0 8.0 5.8 R (21255300011110, sec 1. 7 1. 4 3. 7 3. 2+ 1. 7 1. 8 1. 9 2. 5 2. 5 2. 5+ 2. 5 2. 7 2. 7 2.9 2. 7+ 3.1 2. 6

esu

Contraction, percent 54 26 23 82 35 37 35 46 37- 43 46 49 49 44 43 62 CO disappearance 79 67 54 81 88 67 71 74 83 79 Observed oil, cc./m. 20 19 28 31 44 38 41 101 80 93 80 63 7O 69 v. 50 93 Water, cc./m. 167 81 80 110 165 142 168 129 141 160 121 95 154 Selectivity-- OO-CO2, percent.-- 12. 5 14. 7 22.4 15.4 13.2 20. 5 23. 2 23. 2 11 17.4 17.4 20.0 17.9 (JO-)OH 14.5 17. 8 12. 0 12. 0 7. 4

From the above data it is apparent that the product and results in general compare favorably with fluid-bed operations using a conventional dense phase of catalyst. It should be noted that considerably lower ratio of hydrogen to carbon monoxide in the feed gas is used than is usually practiced in fluid-bed operations. Since the concentration of catalyst is so small as compared to other type operations, the amount of catalyst per unit volume of gas throughput is correspondingly small.

It is assumed that a catalyst slip of about 50 per cent is present in the vertical section of the reaction zone. Actually, therefore, the catalyst loading of pounds of catalyst per cubic foot of gas may be about half the concentration of catalyst in the reaction zone. Thus, for a concentration of 8 pounds per cubic foot of gas, the catalyst flow was calculated to be about 1000 pounds per hour.

When operating at 1.4 ratio of hydrogen to carbon monoxide, as in runs I3 to I5 and [6 to 2|, with a recycle ratio of about 5:1 to 6:1 the gas inlet composition (including fresh feed and recycle gas) was about 18 per cent carbon monoxide and a ratio of HzICO of about 2:1 to 25:1 existed inthe reaction zone.

In certain runs, it was found that with a catalyst loading represented by a differential pressure of 80 inches of water the flow of catalyst in the reaction zone was too great to permit uniform settling in the standpipe. However, with a larger diameter or longer standpipe or bigger slide valve, loadings higher than that represented by 80 inches of Water could be used satisfactorily. In run 24 a combined inlet gas temperature of F. was used. When operating at a combined recycle and fresh feed gas temperature as low as 150 F. it was necessary to maintain the catalyst temperature at least 560 F. to ensure a sufficiently high reaction temperature. The catalyst was maintained at about 560 F. in the standpipe by heating the catalyst electrically, as previously discussed. With the present equipment, the catalyst could not be maintained at a high enough temperature to permit a lower inlet gas temperature. It is believed that an inlet gas temperature as low as 100 F. or lower could be used if the catalyst could be maintained at 600 F. in the standpipe.

In runs Nos. 20 and 2| the selectivity appeared to be good with 21 per cent CO COz and a small amount of CO CH4. Analysis of the product of run 20 indicated that about 40 per cent of the condensed oil comprised oxygenated compounds of which 15 per cent were acids. The water contained about 14 per cent chemicals, approximately half of which were acids. A yield of about 40 cc./m. of total oxygenated compounds was obtained.

Operations at 250 pounds per square inch pressure have proved to be successful with results somewhat similar to those at lower pressures.

As will be noted, no external cooling of the reaction system was used during the present set of runs for the hydrogenation of carbon monoxide by the process of this invention. In fact, heat was added to the system through the catalyst standpipe by electrical heating coils around the jacket of the standpipe which contained Dowtherm under pressure. The reaction zone was heavily lagged and electrically wound to produce substantially adiabatic conditions in that zone. If the entire assembly of apparatus had been completely adiabatic, and if the reaction had been carried out on a larger scale, cooling of the standpipe or the reaction zone would probably have been necessary. Such cooling could have been carried out by contact with a cooling medium surrounding the reaction zone, such as cooling with water or Dowtherm, or by introducing relatively cold feed gas or cold recycle gas or both.

Prior to run 12 an iron catalyst similar to that previously described, but which had been used in conventional dense-phase, fluid-bed operation for about 400 hours, was substituted for the original catalyst in carrying out the process. The composition of this catalyst prior to use in the process is shown in column 1 of Table III. After the catalyst had been used for a period of two days, five days, and nine days, analysis of the catalyst was made to determine the effect of this type of reaction on the carbon and oil con-.

tent of the catalyst, and also upon the catalyst The data of Table III indicates only a slight change in chemical composition and size of the catalyst from the original material introduced at the beginning of run It. The wax content increased about 1.5 per cent and the fixed carbon content increased about 2.3 per cent. A decrease of less than 2 per cent in the total iron was noted. The particle sizeanalysis indicates that the catalyst became slightly finer on use. From these results the catalyst composition in the circulating system or the present invention was probably more stable than with the con U gentional dense-phase operation using a fluid This invention has been described with particular reference to specific apparatus and to specific conditions of operation. However, various modifications of the apparatus and various operating conditions may be used within the limits disclosed without departing from the scope of this invention. Essentially, the apparatus design itself must permit high velocity of gases in at least a portion of the reaction section.

Having described our invention, we claim:

1. A process for the hydrogenation of an oxide of carbon to produce organic compounds having more than one carbon atom. per molecule'which comprises introducing a finely-divided metal hydrogenating catalyst comprising iron and a gaseous mixture comprising hydrogen and an oxide of carbon into the lower portion of an elongated catalyst in gases is smaller than the concentration of the same catalyst when utilized under eperating conditions to obtain a dense phase catalyst bed and such that finely-divided metal catalyst moves in the direction of flow of the gases through said elongated reaction zone, maintaining suitable conditions of temperature, pressure and residence time for the conversion of a major proportion of said oxide of carbon in said elongated reaction zone, maintaining the temperature of said reaction zone at at least 550 F. and within a relatively narrow range of not greater than 100 F. by indirect heat exchange with a cooling medium which is at a temperature substantially lower than the temperature of reaction, passing from the upper portion of said elongated reaction zone a gaseous effluent containing finely-divided metal hydrogenating catalyst through a confined passageway to a separation zone in which finelydivided catalyst is separated from gases, maintaining the pressure dro in the confined passageway between said reaction zone and said separation zone not greater than about inches of water, recovering from said eiiluent organic com pounds having more than one carbon atom per molecule as products of the process, withdrawing catalyst from said separation zone, andpassing catalyst thus separated to the lower portion of said elongated reactionzone at a temperature within about 20 F. of the reaction temperature through a confined passageway.

2. The process of claim 1 in which the termperature oithe cooling medium is at least F. below the temperature of reaction.

3. A process for the hydrogenation of carbon monoxide to produce organic compounds having more than one carbon atom per molecule which comprises introducing a finely-divided, hydrogenating catalyst com-prising iron and an alkali as a promoter and a gaseous mixture comprising hydrogenv and carbon monoxide into the lower portionofan elongated reaction zone, passing gases upward throughsaid reaction zone at a linear velocity of at least about 5 feet per second to suspend said catalyst in the gases such that the concentration of catalyst in said gases is smaller than the concentration of the same catalyst when utilized under operating conditions to obtain a densephase catalyst bed and such that finely-divided catalyst moves in the direction of flow of said gases in said elongated reaction zone, maintaining a temperature of reaction between about 550 and about 650 F. and a pressure be tween about 10. and about 500 pounds 'per square inch gage such that a major proportion of the carbon monoxide is converted, controlling the temperatureof reaction by indirect heat exchange with a cooling medium which is at a substantially lower temperature than said reaction temperature, removing from the upper portion of said elongated reaction zone a gaseous effluent containing finely-divided catalyst and passing same through an unrestricted confined passageway whereby the pressure drop in said passageway is not greater than about 80 inches of water to a separation zone in which'catalyst is separated from gases, recovering organic compounds having more-than one carbon atom per molecule from the gaseous efiiuent as products of the process, with drawing catalyst from said separation zone, and passing catalyst thus separated directly to the lower portion of said elongated reaction zone at a temperature not lower than about 560 F. through a confined passageway. I

4. The process of claim 3 in which the linear gas velocity is between about 8 and about 40 feet per second.

.5. A process for the hydrogenation of carbon monoxide to produce organic compounds having more than one carbon atom per molecule which comp-rises introducing a finely-divided metal hydrogenating catalyst comprising iron and a gaseous mixture comp-rising hydrogen and carbon monoxide into the lower portion of an elongated reaction zone, passing gases upward through said elongated reaction zone at a linear velocity of at least about 6 feet per second to suspend said catalyst in the gases and such that finely-divided catalyst continuously moves in the direction of flow of gases in said elongated reaction zone, maintaining a reaction temperature between about 550 and about 650 F. and a reaction pressure between about 10 and about 500 pounds per square inch gage such that a major proportion of the carbon monoxide is converted, maintaining the reaction temperature within a relatively narrow range within the above temperature range by indirect heat exchange with a cooling medium which is at a temperature substantially below the temperature of reaction, removing from the upper portion of said elongated reaction zone an eiiiuent containing finely-divided catalyst and passing same through a confined passageway to a separation zone in which catalyst is separated from gases, maintaining the pressure drop in said confined passageway between said reaction zone and said separation zone not greater than 80 inches of water, cooling and condensing said effiuent to form. a vapor phase and a liquid phase comprising organic compounds, recovering from said liquid phase organic compounds as products of the process, recycling at least a portion of the vapor phase to the lower portion of said elongated reaction zone in a volumetric ratio of recycle gas to fresh feed gas of about 0.5 1 to about 10:1, withdrawing catalyst from said separation zone, contacting catalyst thus separated with a hydrogen-containing gas, returning catalyst thus separated and contacted with a hydrogen-containing gas to the lower portion of said elongated reaction zone at a temperature not lower than about 560 F., controlling the rate of return of said separated catalyst to the lower portion of said elongated reaction zone such that catalyst loading rate is at least 4 pounds per cubic foot of gas and such that the concentration of catalyst into said elongated reaction zone is less than about 25 pounds per cubic foot.

6. The process of claim 5 in which said hydrogenating catalyst comprising iron contains an alkali and the pressure is between about 80 and about 300 pounds per square inch gage.

'7. The process of claim 5 in which said catalyst concentration is between about 1 and about 18 pounds per cubic foot.

8. A process for the hydrogenation of carbon monoxide to produce organic compounds having more than one carbon atom per molecule which comprises introducing a finely divided hydrogenating catalyst comprising iron and an alkali as its promoter containing no more than a minor proportion by weight of material whose average particle diameter is greater than 250 microns and a gaseous mixture comprising hydrogen and carbon monoxide into the lower portion of an elongated reaction zone, passing gases upwardly through said elongated reaction zone at a linear velocity of at least about 5 feet per second to suspend said finely divided catalyst in the gases such that the concentration of catalyst in said gases is not greater than about 25 pounds per cubic foot of gas and such that finely divided catalyst moves in the direction of flow of said gases in said elongated reaction zone, maintaining a temperature of reaction between about 550 F. and about 650 F. and a pressure between about 10 and about 500 pounds per square inch gage such that a major proportion of the carbon monoxide is converted, controlling the temperature of reaction by indirect heat exchange with a cooling medium which is at a substantially lower temperature than the reaction temperature, removing from the upper portion of said elongated reaction zone a gaseous effluent containing finely divided catalyst and passing same through an unrestricted confined passageway whereby the pressure drop in said passageway is not greater than about 80 inches of water to a separation zone in which catalyst is separated from gases, recovering organic compounds having more than one carbon atom per molecule from the gaseous efiluent as products of the process, withdrawing catalyst from said separation zone, stripping catalyst thus withdrawn with a hydrogen containing gas, and returning catalyst thus separated and stripped directly to the lower portion of said elongated reaction zone at a substantially higher temperature than the gaseous mixture comprising hydrogen and carbon monoxide introduced into said reaction zone and at a temperature of at least 560 F.

9. The process of claim 8 in which the linear velocity of the gases passing upwardly through said elongated reaction zone is between about 8 and about feet per second and the concentration of catalyst in the gases in said reaction zone is between about 3 and about 12 pounds per cubic foot of gas.

HENRY G. MCGRATH. LUTHER R. HILL.

References Cited in the file of this patent UNITED STATES PATENTS Number Name Date 1,799,858 Miller Apr. 7, 1931 2,231,424 Huppke Feb. 11, 1941 2,266,161 Campbell et a1 Dec. 16, 1941 2,325,516 Holt et al July 27, 1943 2,353,505 Scheineman July 11, 1944 2,402,875 Cornell June 25, 1946 2,448,279 Rubin Aug. 31, 1948 2,463,912 Scharmann Mar. 8, 1949 2,464,505 Hemminger Mar. 15, 1949 2,472,377 Keith June 7, 1949 2,481,089 Dickinson H Sept. 6. 1949 FOREIGN PATENTS Number Country Date 533,037 Germany Sept. 8, 1931 

1. A PROCESS FOR THE HYDROGENATION OF AN OXIDE OF CARBON TO PRODUCE ORGANIC COMPOUNDS HAVING MORE THAN ONE CARBON ATOM PER MOLECULE WHICH COMPRISES INTRODUCING A FINELY-DIVIDED METAL HYDROGENATING CATALYST COMPRISING IRON AND A GASEOUS MIXTURE COMPRISING HYDROGEN AND AN OXIDE OF CARBON INTO THE LOWER PORTION OF AN ELONGATED REACTION ZONE, PASSING GASES UPWARD THROUGH SAID ELONGATED REACTION ZONE AT A LINEAR VELOCITY OF AT LEAST ABOUT 5 FEET PER SECOND TO SUSPEND SAID CATALYST IN THE GASES SUCH THAT THE CONCENTRATION OF CATALYST IN GASES IS SMALLER THAN THE CONCENTRATION OF THE SAME CATALYST WHEN UTILIZED UNDER OPEATING CONDITIONS TO OBTAIN A DENSE PHASE CATALYST BED AND SUCH THAT FINELY-DIVIDED METAL CATALYST MOVES IN THE DIRECTION OF FLOW OF THE GASES THROUGH SAID ELONGATED REACTION ZONE, MAINTAINING SUITABLE CONDITIONS OF TEMPERATURE, PRESSURE AND RESIDENCE TIME FOR THE CONVERSION OF A MAJOR PROPORTION OF SAID OXIDE OF CARBON IN SAID ELONGATED REACTION ZONE, MAINTAINING THE TEMPERATURE OF SAID REACTION ZONE AT AT LEAST 550* F. AND WITHIN A RELATIVELY NARROW RANGE OF NOT GREATER THAN 100* F. BY INDIRECT HEAT EXCHANGE WITH A COOLING MEDIUM WHICH IS AT A TEMPERATURE SUBSTANTIALLY LOWER THAN THE TEMPERATURE OF REACTION, PASSING FROM THE UPPER PORTION OF SAID ELONGATED REACTION ZONE A GASEOUS EFFLUENT CONTAINING FINELY-DIVIDED METAL HYDROGENATING CATALYST THROUGH A CONFINED PASSAGEWAY TO A SEPARATION ZONE IN WHICH FINELYDIVIDED CATALYST IS SEPARATED FROM GASES, MAINTAINING THE PRESSURE DROP IN THE CONFINED PASSAGEWAY BETWEEN SAID REACTION ZONE AND SAID SEPARATION ZONE NOT GREATER THAN ABOUT 80 INCHES OF WATER, RECOVERING FROM SAID EFFLUENT ORGANIC COMPOUNDS HAVING MORE THAN ONE CARBON ATOM PER MOLECULE AS PRODUCTS OF THE PROCESS, WITHDRAWING CATALYST FROM SAID SEPARATION ZONE, AND PASSING CATALYST THUS SEPARATED TO THE LOWER PORTION OF SAID ELONGATED REACTION ZONE AT A TEMPERATURE WITHIN ABOUT 20* F. OF THE REACTION TEMPERATURE THROUGH A CONFINED PASSAGEWAY. 